Recovery of butanol from a mixture of butanol, water, and an organic extractant

ABSTRACT

A process for recovering butanol from a mixture comprising a water-immiscible organic extractant, water, butanol, and optionally a non-condensable gas, is provided. The butanol is selected from 1-butanol, 2-butanol, isobutanol, and mixtures thereof. The extractant comprises at least one solvent selected from the group consisting of C 7  to C 22  fatty alcohols, C 7  to C 22  fatty acids, esters of C 7  to C 22  fatty acids, C 7  to C 22  fatty aldehydes, and mixtures thereof.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of priority to U.S. Provisional Patent Application Ser. No. 61/263,502, filed on Nov. 23, 2009, the entirety of which is herein incorporated by reference.

FIELD OF THE INVENTION

Processes for recovering butanol from a butanol-containing organic phase obtained from an extractive fermentation process are provided. Specifically, processes for separating butanol from a mixture comprising butanol, water, a water-immiscible organic extractant, and optionally a non-condensable gas, by distillation and use of an entrainer are provided.

BACKGROUND OF THE INVENTION

Butanol is an important industrial chemical with a variety of applications, such as use as a fuel additive, as a blend component to diesel fuel, as a feedstock chemical in the plastics industry, and as a foodgrade extractant in the food and flavor industry. Each year 10 to 12 billion pounds of butanol are produced by petrochemical means. As the projected demand for butanol increases, interest in producing butanol from renewable resources such as corn, sugar cane, or cellulosic feeds by fermentation is expanding.

In a fermentative process to produce butanol, in situ product removal advantageously reduces butanol inhibition of the microorganism and improves fermentation rates by controlling butanol concentrations in the fermentation broth. Technologies for in situ product removal include stripping, adsorption, pervaporation, membrane solvent extraction, and liquid-liquid extraction. In liquid-liquid extraction, an extractant is contacted with the fermentation broth to partition the butanol between the fermentation broth and the extractant phase. The butanol and the extractant are recovered by a separation process, for example by distillation. In the recovery process, the butanol can also be separated from any water, non-condensable gas, and/or fermentation by-products which may have been removed from the fermentation broth through use of the extractant.

U.S. patent application Ser. No. 12/478,389 filed on Jun. 4, 2009, discloses methods for producing and recovering butanol from a fermentation broth, the methods comprising the step of contacting the fermentation broth with a water immiscible organic extractant selected from the group consisting of C₁₂ to C₂₂ fatty alcohols, C₁₂ to C₂₂ fatty acids, esters of C₁₂ to C₂₂ fatty acids, C₁₂ to C₂₂ fatty aldehydes, and mixtures thereof, to form a two-phase mixture comprising an aqueous phase and a butanol-containing organic phase.

U.S. Provisional Patent Application Nos. 61/168,640; 61/168,642; and 61/168,645; filed concurrently on Apr. 13, 2009; and 61/231,697; 61/231,698; and 61/231,699; filed concurrently on Aug. 6, 2009, disclose methods for producing and recovering butanol from a fermentation medium, the methods comprising the step of contacting the fermentation medium with a water-immiscible organic extractant comprising a first solvent and a second solvent, the first solvent being selected from the group consisting of C₁₂ to C₂₂ fatty alcohols, C₁₂ to C₂₂ fatty acids, esters of C₁₂ to C₂₂ fatty acids, C₁₂ to C₂₂ fatty aldehydes, and mixtures thereof, and the second solvent being selected from the group consisting of C₇ to C₁₁ alcohols, C₇ to C₁₁ carboxylic acids, esters of C₇ to C₁₁ carboxylic acids, C₇ to C₁₁ aldehydes, and mixtures thereof, to form a two-phase mixture comprising an aqueous phase and a butanol-containing organic phase.

U.S. Provisional Patent Application Nos. 61/225,659 and 61/225,662, filed concurrently on Jul. 15, 2009, disclose processes for separating butanol from a mixture comprising butanol, water, a water-immiscible organic extractant, and optionally a non-condensable gas.

Processes for recovering butanol from a butanol-containing extractant phase obtained by in situ product removal from a fermentation broth continue to be sought. Economical processes for recovering butanol substantially free of water and of the extractant are desired. Also desired are separation processes which minimize degradation of the extractant, as well as processes which provide improved efficiency for the desired separations.

SUMMARY OF THE INVENTION

The present invention provides a process for separating a butanol and/or isomers thereof, selected from the group consisting of 1-butanol, 2-butanol, isobutanol, and mixtures thereof, from a feed comprising a water-immiscible organic extractant, water, the butanol, and optionally a non-condensable gas. The separation is made through a combination of distillation, decantation, and use of an entrainer.

In one aspect, the present invention is a process comprising the steps:

-   -   a) introducing a feed comprising:         -   (i) a water-immiscible organic extractant,         -   (ii) water,         -   (iii) at least one isomer of butanol, and         -   (iv) optionally a non-condensable gas     -    into a first distillation column, wherein the first         distillation column comprises a stripping section and optionally         a rectifying section at an introduction point above the         stripping section, the first distillation column having an         operating temperature T₁ and an operating pressure P₁ at a         predetermined point in the stripping section, wherein T₁ and P₁         are selected to produce a first bottoms stream and a first         vaporous overhead stream, the first bottoms stream comprising         the water-immiscible organic extractant and water and being         substantially free of butanol, and the first vaporous overhead         stream comprising water, butanol, optionally the extractant, and         optionally the non-condensable gas;     -   b) introducing a water-immiscible organic entrainer to at least         one appropriate process stream or vessel;     -   c) condensing the first vaporous overhead stream to produce a         gas phase and recover a first mixed condensate, wherein the         first mixed condensate comprises:         -   (i) an organic phase comprising butanol, entrainer, and             water; and         -   (ii) an aqueous phase comprising water and butanol; and     -   wherein the first mixed condensate comprises sufficient         entrainer to provide phase separation of the organic and the         aqueous phases;     -   d) introducing at least a portion of the aqueous phase to the         first distillation column;     -   e) introducing a first portion of the organic phase into a         second distillation column having at least a stripping section;         and     -   f) operating the second distillation column to produce a second         bottoms stream comprising butanol and being substantially free         of water and entrainer, and a second vaporous overhead stream         comprising butanol, entrainer, and water;         wherein

the extractant is selected such that it (A) preferentially extracts butanol over water and (B) is separable from butanol by distillation; and

the entrainer is selected such that it (C) has a higher vapor pressure than butanol and (D) is separable from butanol by distillation.

In embodiments, the isomer of butanol is isobutanol. In embodiments, the butanol in the feed is produced by fermentation of a feedstock such as corn or sugar cane. In embodiments, the feed further comprises ethanol and the second bottoms stream further comprises ethanol. In embodiments, there is a process to process heat exchange between the feed introduced to the first distillation column and the first bottoms stream. In embodiments, the entrainer comprises pentane, hexane, hexene, cyclohexane, benzene, toluene, or xylene. In embodiments, the extractant comprises at least one solvent selected from the group consisting of C₇ to C₂₂ fatty alcohols, C₇ to C₂₂ fatty acids, esters of C₇ to C₂₂ fatty acids, C₇ to C₂₂ fatty aldehydes, C₇ to C₂₂ fatty amides, and mixtures thereof. In embodiments, the extractant comprises at least one solvent selected from the group consisting of C₁₂ to C₂₂ fatty alcohols, C₁₂ to C₂₂ fatty acids, esters of C₁₂ to C₂₂ fatty acids, C₁₂ to C₂₂ fatty aldehydes, C₁₂ to C₂₂ fatty amides, and mixtures thereof. In embodiments, the extractant comprises oleyl alcohol. In embodiments, the isomer of butanol is 1-butanol, 2-butanol, isobutanol, or mixtures thereof.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 illustrates one embodiment of a system useful for practicing the process of the invention.

FIG. 2 illustrates a process schematic diagram used in modeling example embodiments of the process of the invention.

DETAILED DESCRIPTION OF THE INVENTION

Applicants specifically incorporate the entire contents of all cited references in this disclosure. Further, when an amount, concentration, or other value or parameter is given as either a range, preferred range, or a list of upper preferable values and lower preferable values, this is to be understood as specifically disclosing all ranges formed from any pair of any upper range limit or preferred value and any lower range limit or preferred value, regardless of whether ranges are separately disclosed. Where a range of numerical values is recited herein, unless otherwise stated, the range is intended to include the endpoints thereof, and all integers and fractions within the range. It is not intended that the scope of the invention be limited to the specific values recited when defining a range.

The following definitions are used in this disclosure:

Butanol as used herein means 1-butanol (1-BuOH), 2-butanol (2-BuOH), and/or isobutanol (iBuOH or I-BUOH), individually or as mixtures thereof.

“In Situ Product Removal” as used herein means the selective removal of a specific fermentation product from a biological process such as fermentation to control the product concentration in the biological process.

“Fermentation broth” as used herein means the mixture of water, sugars, dissolved solids, suspended solids, microorganisms producing butanol, product butanol and all other constituents of the material held in the fermentation vessel in which product butanol is being made by the reaction of sugars to butanol, water and carbon dioxide (CO₂) by the microorganisms present. The fermentation broth is the aqueous phase in biphasic fermentative extraction. From time to time, as used herein the term “fermentation medium” may be used synonymously with “fermentation broth”.

“Fermentation vessel” as used herein means the vessel in which the fermentation reaction by which product butanol is made from sugars is carried out. The term “fermentor” may be used synonymously herein with “fermentation vessel”.

The term “effective titer” as used herein, refers to the total amount of butanol produced by fermentation per liter of fermentation medium. The total amount of butanol includes: (i) the amount of butanol in the fermentation medium; (ii) the amount of butanol recovered from the organic extractant; and (iii) the amount of butanol recovered from the gas phase, if gas stripping is used.

The term “aqueous phase titer” as used herein, refers to the concentration of butanol in the fermentation broth. Where indicated, the term also refers to the concentration of ethanol in the fermentation broth.

“Stripping” as used herein means the action of transferring all or part of a volatile component from a liquid stream into a gaseous stream.

“Stripping section” as used herein means that part of the contacting device in which the stripping operation takes place.

“Rectifying” as used herein means the action of transferring all or part of a condensable component from a gaseous stream into a liquid stream in order to separate and purify lower boiling point components from higher boiling point components.

“Rectifying section” as used herein means the section of the distillation column above the feed point, i.e. the trays or packing material located above the point in the column where the feed stream enters, where the rectifying operation takes place.

The term “separation” as used herein is synonymous with “recovery” and refers to removing a chemical compound from an initial mixture to obtain the compound in greater purity or at a higher concentration than the purity or concentration of the compound in the initial mixture.

The term “water-immiscible” refers to a chemical component, such as an extractant or solvent, which is incapable of mixing with an aqueous solution, such as a fermentation broth, in such a manner as to form one liquid phase.

The term “extractant” as used herein refers to one or more organic solvents which are used to extract butanol from a fermentation broth.

The term “entrainer” as used herein refers to a third organic component which, when added to an azeotrope formed by a binary mixture, either facilitates or improves the separation of the components of the binary mixture into two liquid phases.

The term “organic phase”, as used herein, refers to the non-aqueous phase of a biphasic mixture obtained by contacting a fermentation broth with a water-immiscible organic extractant.

The term “fatty acid” as used herein refers to a carboxylic acid having a long, aliphatic chain of C₇ to C₂₂ carbon atoms, which is either saturated or unsaturated.

The term “fatty alcohol” as used herein refers to an alcohol having a long, aliphatic chain of C₇ to C₂₂ carbon atoms, which is either saturated or unsaturated.

The term “fatty aldehyde” as used herein refers to an aldehyde having a long, aliphatic chain of C₇ to C₂₂ carbon atoms, which is either saturated or unsaturated.

The term “fatty amide” as used herein refers to an amide having a long, aliphatic chain of C₁₂ to C₂₂ carbon atoms, which is either saturated or unsaturated.

Non-condensable gas means a gas that is not condensed at an operating temperature of the process described herein.

The terms “° C.” and “C” mean degrees Celsius.

The term “deg” means degrees.

The term “g/L” means grams per liter.

The term “ppm” means parts per million.

The term “kg/hr” means kilograms per hour.

The term “atm” means atmosphere.

Butanol-containing extractant streams useful as a feed in the processes of the invention include any organic phase obtained from an extractive fermentation wherein butanol is produced as a fermentation product. Typical butanol-containing extractant streams include those produced in “dry grind” or “wet mill” fermentation processes in which in situ product removal is practiced using liquid-liquid extraction of the fermentation broth with an organic extractant. After extraction, the extractant stream typically comprises butanol, water, and the extractant. The extractant stream may optionally comprise a non-condensable gas, which can be a gas that is inert or otherwise non-reactive with other feed components under the operating conditions of the present invention. Such gases can be selected from gases in the group consisting of, for example, carbon dioxide, nitrogen, hydrogen, noble gases such as argon, or mixtures of any of these. The extractant stream may optionally further comprise fermentation by-products having sufficient solubility to partition into the extractant phase. The extractant stream may optionally contain solids, for example biomass or solids from the fermentation. Butanol-containing extractant streams useful as a feed in the processes of the invention include streams characterized by a butanol concentration in the feed from about 0.1 weight percent to about 40 weight percent, for example from about 2 weight percent to about 40 weight percent, for example from about 5 weight percent to about 35 weight percent, based on the weight of the feed. Depending on the efficiency of the extraction, the aqueous phase titer of butanol in the fermentation broth can be, for example, from about 5 g/L to about 85 g/L, or from about 10 g/L to about 40 g/L. In embodiments, the effective titer of butanol recovered from the process is at least about 40 g/L, at least about 50 g/L, at least about 60 g/L, at least about 70 g/L, at least about 80 g/L, at least about 90 g/L, or at least about 100 g/L.

Butanol-containing extractant streams useful as a feed may further comprise ethanol. Such extractant streams may be characterized by a butanol concentration as described above herein and by an ethanol concentration in the feed from about 0.01 weight percent to about 10 weight percent, for example from about 0.2 weight percent to about 2 weight percent, for example from about 0.5 weight percent to about 1 weight percent, based on the weight of the feed. Depending on the efficiency of the extraction, the aqueous phase titer of ethanol in the fermentation broth can be, for example, from about 0.1 g/L to about 20 g/L, or from about 1 g/L to about 5 g/L. The ethanol may be obtained in the fermentation broth as a by-product from recombinant butanol-producing microorganisms, for example.

The extractant is a water-immiscible organic solvent or solvent mixture having characteristics which render it useful for the extraction of butanol from a fermentation broth. The extractant preferentially partitions butanol from the aqueous phase, for example by at least a 1.1:1 concentration ratio, such that the concentration of butanol in the extractant phase is at least 1.1 times that in the aqueous phase when evaluated in a room-temperature extraction of an aqueous solution of butanol. In embodiments, the extractant preferentially partitions butanol from the aqueous phase by at least a 2:1 concentration ratio, such that the concentration of butanol in the extractant phase is at least two times that in the aqueous phase when evaluated in a room-temperature extraction of an aqueous solution of butanol. In embodiments, the extractant preferentially partitions butanol from the aqueous phase by at least a 3:1 concentration ratio, by at least a 4:1 concentration ratio, by at least a 5:1 concentration ratio, by at least a 6:1 concentration ratio, by at least a 8:1 concentration ratio, by at least a 10:1 concentration ratio or by at least a 20:1 concentration ratio.

To be of practical use in the butanol recovery process, the extractant is separable from butanol by distillation, having a boiling point at atmospheric pressure which is at least about 30 degrees Celsius higher than that of the butanol to be recovered, or for example at least about 40 degrees higher, or for example at least about 50 degrees higher. A mixture of higher boiling extractants is expected to behave in a fundamentally similar way to a single extractant provided that the boiling point of the mixture, or the boiling point of the lowest boiling solvent of the mixture, is significantly higher than the boiling points of water and butanol, for example at least about 30 degrees higher.

The extractant can comprise at least one solvent selected from the group consisting of C₇ to C₂₂ fatty alcohols, C₇ to C₂₂ fatty acids, esters of C₇ to C₂₂ fatty acids, C₇ to C₂₂ fatty aldehydes, C₇ to C₂₂ fatty amides, and mixtures thereof. The extractant can comprise at least one solvent selected from the group consisting of C₁₂ to C₂₂ fatty alcohols, C₁₂ to C₂₂ fatty acids, esters of C₁₂ to C₂₂ fatty acids, C₁₂ to C₂₂ fatty aldehydes, C₁₂ to C₂₂ fatty amides, and mixtures thereof. Examples of suitable extractants include an extractant comprising at least one solvent selected from the group consisting of oleyl alcohol, behenyl alcohol, cetyl alcohol, lauryl alcohol, myristyl alcohol, stearyl alcohol, oleic acid, lauric acid, myristic acid, stearic acid, methyl myristate, methyl oleate, lauric aldehyde, 1-nonanol, 1-decanol, 1-undecanol, 2-undecanol, 1-nonanal, 2-butyloctanol, 2-butyl-octanoic acid and mixtures thereof. In embodiments, the extractant comprises oleyl alcohol. In embodiments, the extractant comprises a branched chain saturated alcohol, for example, 2-butyloctanol, commercially available as ISOFAL® 12 (Sasol, Houston, Tex.) or Jarcol I-12 (Jarchem Industries, Inc., Newark, N.J.). In embodiments, the extractant comprises a branched chain carboxylic acid, for example, 2-butyl-octanoic acid, 2-hexyl-decanoic acid, or 2-decyl-tetradecanoic acid, commercially available as ISOCARB® 12, ISOCARB® 16, and ISOCARB® 24, respectively (Sasol, Houston, Tex.).

Such organic extractants can be available commercially from various sources, such as Sigma-Aldrich (St. Louis, Mo.), in various grades, many of which may be suitable for use in extractive fermentation to produce or recover butanol. Technical grades contain a mixture of compounds, including the desired component and higher and lower fatty components. For example, one commercially available technical grade oleyl alcohol contains about 65% oleyl alcohol and a mixture of higher and lower fatty alcohols.

The invention provides processes for separating or recovering butanol from a feed comprising a water-immiscible organic extractant, water, at least one isomer of butanol, and optionally a non-condensable gas. Separation of the butanol from the feed is achieved through a combination of distillation, decantation, and the use of an entrainer. The distillation involves the use of at least two distillation columns. The first column, in combination with the entrainer and decantation, effects a separation of any non-condensable gas, such as carbon dioxide, and butanol from the extractant, for example oleyl alcohol, and water. The entrainer is added to an appropriate process stream or vessel in an amount such that the first mixed condensate contains sufficient entrainer to provide phase separation of the organic and the aqueous phases. Such phase separation typically occurs in a decanter. By “phase separation” is meant the physical formation of two liquid phases, one mostly aqueous and one mostly organic, from one initial liquid phase containing water and organics. The physical characteristics of the entrainer and its concentration in the initial liquid phase, as well as the concentrations of butanol and ethanol in the initial liquid phase, determine if phase separation occurs under the selected process conditions. Temperature and pressure, for example, can also affect phase separation. At least a portion of the aqueous phase is returned to the first column; a portion of the organic phase is also returned to the first column. The second column effects a separation of butanol and water and provides butanol as a bottoms stream which is substantially free of water and entrainer. By “substantially free of water” it is meant that less than about 0.01 weight percent of water is present in the bottoms stream. By “substantially free of entrainer” it is meant that less than about 0.01 weight percent of entrainer is present in the bottoms stream.

The entrainer is a water-immiscible organic compound having characteristics which render it useful in the processes of the invention. The entrainer has a sufficiently higher vapor pressure and is more volatile than the butanol isomer to be separated (or than the most volatile butanol isomer in a mixture of butanol isomers) to enable its use in the distillations described herein. For example, when the operating conditions of the first and/or second distillation columns include using about atmospheric pressure at the tops of the columns, the difference in vapor pressure between the entrainer and the most volatile butanol isomer may be about 5 to about 50 psi. When the operating conditions include using less than atmospheric pressure at the tops of the distillation columns, the difference in vapor pressure may be smaller, for example about 10 to about 30 psi. Using an entrainer which is too volatile in relation to the butanol can result in excessive entrainer losses during separation or require much colder chilling media to condense and recover the entrainer. To be suitable for use, the entrainer also has a low molar latent heat, is thermally stable under the operating conditions of the process, and is inert or otherwise non-reactive with other components in the feed stream.

To be of practical use in the butanol recovery process, the entrainer is separable from butanol by distillation and has a boiling point at atmospheric pressure which is lower than that of the butanol to be recovered.

The entrainer can comprise at least one hydrocarbon. The entrainer can be, for example, a saturated or unsaturated, substituted or unsubstituted, aliphatic hydrocarbon. The entrainer can be a substituted or unsubstituted aromatic hydrocarbon. For example, the entrainer may comprise at least one hydrocarbon selected from the group consisting of pentane, hexane, hexene, cyclohexane, benzene, toluene, and xylene. Preferably, the entrainer comprises hexane.

The processes of the invention can be understood by reference to FIG. 1, which illustrates one embodiment of a system useful for practicing the process of the invention. The feed stream 430, obtained from a fermentation vessel (not shown) or an extractor (not shown) in a process for fermentative extraction, is introduced into a first distillation column 650, which has a stripping section and optionally a rectifying section, at a feed point above the stripping section. The feed stream 430 is distilled to provide a first bottoms stream 410 and a first vaporous overhead stream 300 comprising water, butanol, entrainer, and any non-condensable gas present in the feed. An operating temperature T₁ and an operating pressure P₁ at a predetermined point in the stripping section of column 650 are selected so as to provide the first bottoms stream 410 comprising the extractant and water and being substantially free of butanol and substantially free of entrainer. By “substantially free of butanol” it is meant that butanol comprises no more than 0.01 wt % of the bottoms 410. By “substantially free of entrainer” it is meant that the entrainer comprises no more than 0.01 wt % of the bottoms 410. The distillation column 650 can be any conventional column having at least a feed inlet, an overhead vapor outlet, a bottoms stream outlet, a heating means, and a sufficient number of stages to effect the separation of the butanol from the extractant. A rectification section is required when minimum extractant loss in stream 500 is desired and may or may not be combined with the use of an organic reflux stream 490. In the case where the extractant comprises oleyl alcohol, distillation column 650 should have at least 5 stages including a re-boiler.

The first bottoms stream 410 can include from about 3 to about 12 weight percent water, less than about 0.01 weight percent butanol, and less than about 0.01 weight percent entrainer. To ensure that the bottom stream 410 is substantially free of butanol and entrainer, the ratio of the aqueous to organic reflux flows to distillation column 650 should be chosen such that the aqueous reflux (stream 480) exceeds the organic reflux (stream 490) by a ratio higher than the ratio of water to the organic composition of the first mixed condensate. The process may further comprise introducing bottoms stream 410 from the first distillation column into a fermentation vessel (not shown). Alternatively, bottoms stream 410 may be separated (not shown) to obtain a bottoms aqueous phase comprising water and a bottoms organic phase comprising the extractant, introducing at least a portion of the bottoms organic phase into a fermentation vessel, and optionally introducing at least a portion of the bottoms aqueous phase into the same or a different fermentation vessel. The separation may be done, for example, by cooling the bottoms stream 410 until phase separation occurs. These options provide means to recycle the first bottoms stream 410 from the butanol recovery process to the extractive fermentation process.

Optionally, additional stream 435 comprising water, steam, or a mixture thereof may be introduced into the first distillation column 650 at a point anywhere along the column. If water is used, it is preferred that the water be fed together with aqueous stream 480, which is returned as liquid reflux. If steam is used, it is preferred that the steam be fed in the stripping section or from the bottom of the column. The feed point of stream 435 may be the same as or different from the feed point of feed stream 430. The total aqueous return to the column is the sum of the aqueous stream 480 and the optional stream 435, and the total aqueous return to the column should be chosen so as to be sufficient to maintain liquid water throughout all the column trays. The amount of any added water, steam, or a mixture thereof should also be chosen such that, in combination with aqueous stream 480, the total aqueous return to the column exceeds the butanol stream 490 returned to the column by a ratio that is greater than the ratio of the aqueous composition to the organic composition of the first mixed condensate.

The vaporous overhead stream 300 from the first distillation column can include up to about 66 weight percent butanol and from about 20 to about 50 weight percent water. Overhead stream 300 can also include about 5 to about 20 weight percent entrainer. When more entrainer is used in the process than the minimal amount sufficient to provide phase separation of the first mixed condensate under the selected operating conditions, the amount of entrainer in stream 300 is proportionately increased, and the weight percentages of butanol and water are proportionately decreased. The overhead stream includes non-condensable gas that may have been present in the feed. Stream 300 is condensed in a condenser 670 to produce a first mixed condensate stream 350 comprising condensed liquid organics and condensed liquid water. The first mixed condensate stream 350 should comprise sufficient entrainer to provide phase separation of the organic and the aqueous phases. Stream 350 also includes any non-condensable gas present in the feed. The condenser 670 may be of any conventional design.

The mixed condensate stream 350 is introduced into a decanter 700 and allowed to separate into a liquid organic phase and a liquid aqueous phase. The temperature of the decanter is preferably maintained at or below about 40° C. to reduce the amount of butanol, entrainer, and water being stripped out by the non-condensable gas. The liquid organic phase, the lighter liquid phase (the top liquid phase), can include less than about 10 weight percent water, or from about 0.1 to about 5 weight percent water and may further comprise less than about 0.1 weight percent of any extractant which comes overhead in column 650. The liquid organic phase can also include less than about 70 weight percent entrainer, or from about 40 to about 50 weight percent entrainer. When more entrainer is used in the process than the minimal amount sufficient to provide phase separation of the first mixed condensate under the selected operating conditions, the amount of entrainer in the liquid organic phase is proportionately increased, and the weight percentages of butanol and water are proportionately decreased. The residual extractant in the organic phase can be minimized by use of a rectification section in column 650. The liquid aqueous phase includes less than about 30 weight percent, or from about 3 to about 20 weight percent butanol. The liquid aqueous phase can include less than about 5 weight percent entrainer, or less than about 1 weight percent entrainer. The decanter may be of any conventional design.

When a non-condensable gas such as carbon dioxide is present in the feed, the non-condensable gas is present in stream 300 and in stream 350. The process may further comprise the step of purging at least a portion of the gas phase comprising the non-condensable gas from the process, as shown in FIG. 1, in which purge stream 460 comprising the non-condensable gas is shown leaving the decanter 700. Purge stream 460 can further comprise entrainer. To minimize the amount of entrainer which is lost through this purge stream, stream 460 can be partially condensed to recover a portion of the entrainer contained therein, and the recovered entrainer can be returned to the process (not shown). In one embodiment, the process can further comprise the step of partially condensing the gas phase to recover at least a portion of the entrainer and optionally introducing the recovered entrainer to the mixed condensate.

From the decanter 700, at least a portion of the aqueous phase 480 is introduced to the first distillation column 650. The aqueous phase 480 may be introduced as reflux to the column and will typically include all of the aqueous phase separated in the decanter. Introducing stream 480 into column 650 lowers the column temperature and ensures that water is present in the bottoms stream. This is advantageous because higher column temperatures may result in degradation of the extractant in the column bottoms, particularly in the case where fermentation byproducts such as organic acids are present and the extractant contains functional groups, such as unsaturated carbon-carbon bonds in the case of oleyl alcohol, or unfermented sugar and starch. Degradation of any sugar and starch should be avoided as it can cause fouling or plugging in heat exchanger or tray decks. Degradation of the extractant should be avoided as it can cause decreased efficiency in the extractive fermentation process.

The process may optionally further comprise introducing at least a portion of the aqueous phase from the decanter into a fermentation vessel (not shown). This can provide a means to recycle some of the water from the butanol recovery process back to the extractive fermentation process. However, it is normally preferred to recycle water to the fermenter via stream 410, in which the butanol content is lower.

The organic phase 470 leaving the decanter is split into two portions. A first portion of the organic phase, stream 500, is introduced into a second distillation column 950, which has a stripping section, at a feed point above the stripping section. The stream 500 is distilled to provide a second bottoms stream 600 comprising butanol and a second vaporous overhead stream 210 comprising butanol, water, and entrainer. The second distillation column is operated so as to provide the bottoms stream 600 substantially free of water and substantially free of entrainer. By “substantially free of water” it is meant that the bottoms 600 include less than about 0.01 weight percent water. By “substantially free of entrainer” is meant that the bottoms 600 include less than about 0.01 weight percent entrainer. The distillation column 950 can be any conventional column having at least a feed inlet, an overhead vapor outlet, a bottoms stream outlet, a heating means, a stripping section, and a sufficient number of stages to effect the desired separation. Column 950 should have at least 6 stages including re-boiler and can be without a re-boiler when direct steam is injected from the column bottom.

A second portion of the organic phase, stream 490, is introduced into the first distillation column 650. Stream 490 may be introduced as reflux to the column. Introducing stream 490 into column 650 as reflux will suppress extractant loss in vaporous stream 300 of column 650. The ratio of stream 490 over stream 470 can range from about 0.1 to 50 weight percent.

The vaporous overhead stream 210 from the second distillation column 950 can include up to about 20 weight percent butanol and about 5 weight percent water. Stream 210 can further comprise about 50 to about 90 weight percent entrainer. When more entrainer is used in the process than the minimal amount sufficient to provide phase separation of the first mixed condensate under the selected operating conditions, the amount of entrainer in stream 210 is proportionately increased, and the weight percentages of butanol and water are proportionately decreased. Stream 210 is condensed in a condenser 975 to produce a second condensate stream 250 comprising butanol, water, and entrainer. The condenser 975 may be of any conventional design. At least a portion of the second condensate stream 250 may be introduced into the first mixed condensate stream, for example by feeding the second condensate steam 250 into decanter 700. The combination of the first mixed condensate stream and the second condensate stream should contain sufficient entrainer to provide phase separation of the combination into liquid organic and liquid aqueous phases. Any non-condensable gas can be purged as previously described hereinabove.

The vaporous overhead stream 210 may further comprise volatile fermentation byproducts such as acetaldehyde. Optionally, at least a portion of stream 210 may be purged from the process (not shown) to remove volatile fermentation byproducts from the butanol recovery process. Such a purge stream can represent a potential loss of entrainer, which can be compensated for by the introduction of additional entrainer to the process.

In one embodiment, feed stream 430 further comprises ethanol. As described herein above, the feed stream 430 is distilled in the first distillation column 650 to provide the first bottoms stream 410, which comprises the extractant and water and is substantially free of butanol, entrainer, and ethanol, and the first vaporous overhead stream 300, which comprises water, butanol, entrainer, ethanol, and any non-condensable gas present in the feed. The first bottoms stream 410 can include from about 3 to about 15 weight percent water, less than about 0.01 weight percent butanol, less than about 0.01 weight percent ethanol, and less than about 0.01 weight percent entrainer. Feed stream 430 comprising butanol and ethanol is introduced into column 650 at a feed point above the stripping section.

When the feed comprises butanol and ethanol, the vaporous overhead stream 300 from the first distillation column can include 10 to about 40 weight percent butanol, about 1 to about 10 weight percent ethanol, about 10 to about 50 weight percent water, and about 5 to about 30 weight percent entrainer. The composition of stream 300 will depend on the composition of the feed stream, the operating conditions of the first column (for example temperature and pressure), and the amount of entrainer circulating in the process. When more entrainer is used in the process than about the minimal amount sufficient to provide phase separation of the first mixed condensate under the selected operating conditions, the amount of entrainer in stream 300 is proportionately increased, and the weight percentages of butanol, ethanol, and water are proportionately decreased. Condensation of stream 300 produces the first mixed condensate stream 350, which should comprise sufficient entrainer to provide phase separation of the organic and the aqueous phases. The liquid organic phase comprises butanol, entrainer, ethanol, and optionally the extractant. The liquid organic phase can contain less than about 80 weight percent, or from about 40 to about 70 weight percent, entrainer. The liquid aqueous phase is substantially free of entrainer and comprises water, butanol, and ethanol. The liquid aqueous phase can contain less than about 0.1 weight percent entrainer.

As disclosed herein above, the first portion of the organic phase leaving the decanter, stream 500, is distilled in the second distillation column 950 to provide the second bottoms stream 600 and the second vaporous overhead stream 210. When the feed stream 430 comprises butanol and ethanol, the second bottoms stream 600 also comprises butanol and ethanol. In general, operating conditions for the first and second columns can be selected to maintain about the same ratio (on a mass basis) of ethanol to butanol in stream 600 as in feed stream 430. Stream 600 may contain from about 70 weight percent to about 99 weight percent butanol and from about 1 weight percent to about 30 weight percent ethanol. Stream 600 may also contain small quantities of extractant and entrainer. The vaporous overhead stream 210 can contain from about 1 weight percent to about 10 weight percent butanol, from about 1 weight percent to about 10 weight percent ethanol, and from about 70 weight percent to about 90 weight percent entrainer.

The entrainer can be introduced to any appropriate point or points in the portion of the process through which the entrainer circulates. Appropriate process streams or vessels to which the entrainer (or a stream comprising the entrainer) can be added include the feed stream 430, the first vaporous overhead stream 300, the first mixed condensate stream 350, the second condensate stream 250, the portion of the organic phase 500 which is introduced into the second column, the vaporous overhead stream 210 from the second distillation column, the first distillation column 650, the decanter 700, the upper section of the second distillation column 950, and combinations thereof. Shown in FIG. 1 is addition of the entrainer as stream 120 to the decanter 700. When the stream to which the entrainer is introduced is vaporous, the entrainer is preferably preheated and added as a vaporous stream.

Addition of the entrainer to the process can be performed in a continuous manner or in a discontinuous manner, so long as the amount of entrainer in the first mixed condensate, and/or in the combination of the first mixed condensate stream and the second condensate stream, is sufficient to provide phase separation of the organic and the aqueous phases under the operating conditions of the process. An amount of entrainer in excess of that required for phase separation can be used but as this can lead to increased volumes of the process streams comprising the entrainer, increased energy consumption, proportionally larger loss of entrainer from the process, and increased operating cost, the use of significantly excess entrainer in the process is typically not desirable. Use of about the minimal amount of entrainer which is sufficient to provide phase separation under the selected operating conditions is preferred. Typically, make-up entrainer is added to the process to compensate for the entrainer losses which can be incurred when vaporous streams exit the process.

The present processes for separating or recovering butanol provide butanol known to have an energy content similar to that of gasoline and which can be blended with any fossil fuel. Butanol is favored as a fuel or fuel additive as it yields only CO₂ and little or no SO_(X) or NO_(X) when burned in the standard internal combustion engine. Additionally, butanol is less corrosive than ethanol, the most preferred fuel additive to date.

In addition to its utility as a biofuel or fuel additive, the butanol recovered according to the present processes has the potential of impacting hydrogen distribution problems in the emerging fuel cell industry. Fuel cells today are plagued by safety concerns associated with hydrogen transport and distribution. Butanol can be easily reformed for its hydrogen content and can be distributed through existing gas stations in the purity required for either fuel cells or vehicles. Furthermore, the present processes recover butanol obtained from plant derived carbon sources, avoiding the negative environmental impact associated with standard petrochemical processes for butanol production.

One advantage of the present processes for separation or recovery of butanol is that by returning a portion of the aqueous phase from the decanter to the first column, the temperature in the first column is kept relatively low, for example below about 140° C. under any conditions, and in the case of operation at atmospheric pressure, closer to 100° C. Lower temperatures avoid or reduce fouling of the heat exchangers associated with the column, as can occur when the extractant in the bottoms stream degrades, for example through reactions with or catalyzed by fermentation byproducts contained in the bottoms stream. Lower column temperatures also make the recovery process more economical.

An additional advantage is that the first bottoms stream comprising the extractant is substantially free of the butanol product, which contributes to high yield in the recovery process. Being substantially free of butanol also enables optional recycling of the first bottoms stream to the fermentative process. It also simplifies its disposition, should it not be recycled.

Another advantage of the present invention is that it provides a more efficient process due to the increased relative volatility of water, which leads to less boilup being required to operate the second column compared to the case without an entrainer.

Although particular embodiments of the present invention have been described in the foregoing description, it will be understood by those skilled in the art that the invention is capable of numerous modifications, substitutions, and rearrangements without departing from the spirit of essential attributes of the invention. Reference should be made to the appended claims, rather than to the foregoing specification, as indicating the scope of the invention.

The process of the invention can be demonstrated using a computational model of the process. Process modeling is an established methodology used by engineers to simulate complex chemical processes. Process modeling software performs many fundamental engineering calculations, for example mass and energy balances, vapor/liquid equilibrium and reaction rate computations. The modeling of distillation columns is particularity well established. Calculations based on experimentally determined binary vapor/liquid equilibrium and liquid/liquid equilibrium data can predict reliably the behavior of multi-component mixtures. This capability has been expanded to allow modeling of complex multi-stage, multi-component distillation columns using rigorous algorithms like the “inside-out” algorithm developed by Joseph Boston of Aspentech, Inc. of Burlington, Mass. Commercial modeling software, such as Aspen Plus® from Aspentech, can be used in conjunction with physical property databases, such as DIPPR, available from the American Institute of Chemical Engineers, Inc., of New York, N.Y., to develop accurate models and assessments of processes.

EXAMPLES

Examples 1 and 2 were obtained through process modeling using 2-butanol as the butanol isomer, oleyl alcohol as the extractant, and n-hexane as the entrainer. Examples 3 and 4 were obtained through process modeling using isobutanol as the butanol isomer, oleyl alcohol as the extractant, and n-hexane as the entrainer. A small amount of ethanol was included in the feed stream for Examples 3 and 4.

Similar results would be expected for the analogous cases where 1-butanol or mixtures of 1-butanol, 2-butanol, and/or isobutanol were selected as the butanol isomer.

Table 1 lists typical feed compositions of the rich solvent stream, obtained from extractive fermentation, entering the alcohol product recovery area. These compositions were used in modeling the processes of the invention. In the Examples, the term “rich solvent stream” is synonymous with the term “feed stream” used above.

TABLE 1 Feed Compositions (in Weight Percent) of the Rich Solvent Stream from the Extractor Feed Compositions Example 1 Example 2 Example 3 Example 4 iso-butanol — — 5.4% 5.375% 2-butanol 5.40% 13.84% — — ethanol — — 0.54%  0.9954%  Water 7.20%  9.58% 7.2% 7.164% Carbon dioxide 0.20%  0.75% 0.2% 0.199% Oleyl alcohol 87.2% 75.83% 86.66%  86.27%

These composition values for the rich solvent stream were established by a simulation of a dry grind facility using extractive in situ product removal technology producing 50 MM gal/year of 2-butanol or isobutanol, and fermenter broth aqueous phase titers of 10 and 40 g/L 2-butanol or 10 g/L isobutanol respectively. It was assumed that the rich solvent stream was at equilibrium with the fermentation broth and that the solvent flow rate was sufficient to meet the specified annual capacity.

The parameters inputted for the simulations of the embodiments of the processes of the invention are listed in Table 2 and follow a process schematic diagram as shown in FIG. 2. In FIG. 2, “EM10” refers to a heat stream representing process to process heat exchange between the solvent column feed and bottom product via heat exchangers. Block 80 represents a mixer combining the two streams 12 and 19. Block 70 represents a mixer combining the two streams 30 and 21. Block 75 represents a modeling artifact where all entrainer is taken out of the process so that the correct amount of entrainer can be added to the decanter. Certain dimensions and duty results calculated from the process model are also listed in Table 2. These parameters do not include physical property parameters, and those related to convergence and other computational options or diagnostics. The organic reflux to the solvent column is expressed in terms of the split fraction on the total flow of the organic phase 47 from the decanter.

TABLE 2 Conditions Used for Modeling Processes of the Invention Blocks Inputs Example 1 Example 2 Example 3 Example 4 Units Solvent Number of theoretical 15 15 15 15 stages Column stages including re-boiler (65) Column top pressure 1 1 1 1 bar Column bottom pressure 1.1 1.1 1.1 1.1 bar Column internal diameter 4.46 2.83 4.44 4.39 m Column re-boiler duty 92577 48300 99716 101471 MJ/hr Preheated rich solvent 3 3 3 3 stage feed (43) location Organic reflux from 1 1 1 1 stage decanter (48) location Mass fraction water in 1 1 1 1 ppm bottom stream (49) Reflux stream temperature 40 40 40 40 deg C. Preheated rich solvent 330000 131500 330000 330000 kg/hr stream (42) flow rate Preheated rich solvent 85.8 76.8 98.7 98.3 deg C. stream (42) temperature BuOH Number of theoretical 15 15 15 15 stages Column stages including re-boiler (95) Column top pressure 1 1 0.1 0.1 bar Column bottom pressure 1.05 1.05 0.105 0.105 bar Column internal diameter 3.16 3.16 2.81 2.76 m Column re-boiler duty 17366 17396 7494 6061 MJ/hr Organic feed from solvent 1 1 1 1 stage column (50) location Organic feed from solvent 40 40 40 40 deg C. column (50) temperature Entrainer mass fraction in 1 1 1 1 ppm bottom product (60) Decanter Decanter pressure 1 1 1 1 atm (85) Decanter temperature 35 35 35 35 deg C. Entrainer circulation rate 28000 28000 28000 28000 kg/hr Orgsplit Ratio of stream (49) to all 0.2 0.2 0.2 0.2 (90) organic from decanter (47)

Four cases were run to demonstrate the operating requirements of the processes of the invention. Examples 1 and 2 were run to demonstrate the separation of 2-butanol from the rich solvent stream. Examples 3 and 4 were run to demonstrate the separation of isobutanol and a small amount of ethanol from the rich solvent stream. For each case, a particular modification was made to the rich solvent feed flow and compositions from the extractive fermentation process where specific aqueous phase titers were maintained. In each of the independent simulations, column traffic and heat exchanger duties will change because of the feed composition change. By comparing the resulting capital investment and operating costs between different cases, the impact of the rich solvent feed flow and composition on product recovery area performance was quantified. These four examples, however, should not be regarded as process operating limits of this invention.

The term “Solvent Column” is synonymous with the term “first distillation column” used above. The term “BUOH column” is synonymous with the term “second distillation column” used above. The abbreviation “OLEYLOH” refers to oleyl alcohol. The abbreviation “N-C6” refers to n-hexane.

Stream results for Example 1 are listed in Table 3. BUOH column traffic and liquid mass composition profiles are listed in Table 4. Solvent column traffic and liquid mass composition profiles are listed in Table 5.

Stream results for Example 2 are listed in Table 6. BUOH column traffic and liquid mass composition profiles are listed in Table 7. Solvent column traffic and liquid mass composition profiles are listed in Table 8.

Stream results for Example 3 are listed in Table 9. BUOH column traffic and liquid mass composition profiles are listed in Table 10. Solvent column traffic and liquid mass composition profiles are listed in Table 11.

Stream results for Example 4 are listed in Table 12. BUOH column traffic and liquid mass composition profiles are listed in Table 13. Solvent column traffic and liquid mass composition profiles are listed in Table 14.

Other process parameters include the following: 1) the total number of theoretical stages in the solvent column and the feed location of the preheated rich solvent stream; 2) the split fraction of organic reflux to solvent column; 3) the degree of preheating of the rich solvent stream before feeding it to the solvent column; and 4) the amount of entrainer, water, and solvent allowed in the final product. These parameters can be manipulated to achieve optimum separation performance.

Example 1

In this Example, 330,000 kg/hr rich solvent feed 40 containing 5.4 weight percent 2-butanol is heated from 32 to 85.8° C. by a process to process heat exchanger and the resulting stream 41 is fed to the solvent column at stage 3. This feed point divides the solvent column into the rectifying and stripping sections. This rich solvent feed condition corresponds to 10 g/L aqueous phase titer in the fermenter which is maintained during the extractive fermentation process. The separation is realized by a larger diameter solvent column, and higher solvent column re-boiler and condenser duties. Stream 60 is 99.94 weight percent 2-butanol.

Example 2

In this Example, 131,500 kg/hr rich solvent feed 40 containing 13.84 weight percent 2-butanol is heated from 32 to 76.8° C. by a process to process heat exchanger and the resulting stream 41 is fed to the solvent column at stage 3. This feed point divides the solvent column into the rectifying and stripping sections. This rich solvent feed condition corresponds to 40 g/liter aqueous phase titer in the fermenter which is maintained during the extractive fermentation process. The separation is realized by a smaller diameter solvent column, and lower solvent column re-boiler and condenser duties. Stream 60 is 99.99 weight percent 2-butanol.

Example 3

In this Example, 330,000 kg/hr rich solvent feed 40 containing 5.4 weight percent isobutanol and 0.54 weight percent ethanol is heated from 32 to 85.8° C. by a process to process heat exchanger and the resulting stream 41 is fed to the solvent column at stage 3. This feed point divides the solvent column into the rectifying and stripping sections. This rich solvent feed condition is corresponding to 10 g/liter aqueous phase isobutanol titer in the fermenter which is maintained during the extractive fermentation process. In addition, a small amount of ethanol is assumed to be present in the fermentation broth. The mass ratio of ethanol to isobutanol in rich solvent stream is assumed to be 10 wt % in this Example. The separation is realized by a similar diameter solvent column, and a similar solvent column re-boiler and condenser duties as those of Example 1. Stream 60 is 8.9 weight percent ethanol and 90.1 weight percent isobutanol.

Example 4

In this Example, 330,000 kg/hr rich solvent feed (40) is heated from 32 to 85.4° C. by a process to process heat exchanger and the resulting stream (41) is fed to the solvent column at stage 3. This feed point divides the solvent column into the rectifying and stripping sections. This rich solvent feed condition is corresponding to 10 g/liter aqueous phase isobutanol titer in the fermenter which is maintained during the extractive fermentation process. In addition, a small amount of ethanol is assumed to be present in the fermentation broth. The mass ratio of ethanol to isobutanol in rich solvent stream is assumed to be 18.5 wt % in this example. The separation is realized by a similar diameter solvent column, and a similar solvent column re-boiler and condenser duties as those of Example 1. Stream 60 is 15.2 weight percent ethanol and 82.4 weight percent isobutanol.

TABLE 3 Simulated Stream Outputs for Example 1. 11 12 19 20 21 30 33 40 41 Temperature C. 87.6 25 87.6 77.8 64.9 91.6 87.6 32 99.5 Pressure atm 0.99 2 0.99 0.99 0.99 0.99 0.99 1.04 1.09 Vapor Frac 1 0 0.664 0.589 1 1 0.992 0 0 Mole Flow kmol/hr 318.828 324.932 1860.063 2184.987 354.492 1824.399 2178.891 2646.105 2389.797 Mass Flow kg/hr 27474.046 28000 59185.213 87185.081 27021.318 59637.942 86659.259 330000 311516.83 Volume Flow l/hr   9.22E+06 42683.8   3.65E+07   3.67E+07   9.60E+06   5.46E+07   6.38E+07 390661.04 389206.37 Enthalpy MMBtu/hr −4.76E+01 −61.151 −4.70E+02 −5.31E+02 −6.51E+01 −4.29E+02 −4.94E+02 −986.943 −857.501 Mass Flow kg/hr 2-BUOH 33062.959 33062.971 4066.919 28996.04 33062.959 17820 31.152 OLEYLOH 14.35 14.35 <0.001 14.35 14.35 287760 287748.52 WATER 25033.771 25033.627 682.565 24351.207 25033.771 23760 23737.156 N—C6 27474.046 28000 28000 21976.937 5497.109 27474.046 trace CO2 1074.133 1074.134 294.897 779.236 1074.133 660 trace Mass Frac 2-BUOH 0.559 0.379 0.151 0.486 0.382 0.054 100 PPM OLEYLOH 242 PPM 165 PPM trace 241 PPM 166 PPM 0.872 0.924 WATER 0.423 0.287 0.025 0.408 0.289 0.072 0.076 N—C6 1 1 0.321 0.813 0.092 0.317 trace CO2 0.018 0.012 0.011 0.013 0.012 0.002 trace 42 43 46 47 48 49 50 60 Temperature C. 42 85.8 35 35 35 35 35 100.586873 Pressure atm 1.09 1.05 1 1 1 1 1 1.03626943 Vapor Frac 0 0.003 1 0 0 0 0 0 Mole Flow kmol/hr 2389.797 2646.105 23.013 742.353 1419.62 148.471 593.883 239.390326 Mass Flow kg/hr 311516.83 330000 1257.07 55966.55 29961.461 11193.31 44773.24 17751.9221 Volume Flow l/hr 370205.91 608462.5 573446.9 78326.768 31513.163 15665.353 62661.414 24530.4775 Enthalpy MMBtu/hr −896.741 −947.704 −7.017 −194.042 −387.451 −38.808 −155.234 −73.658739 Mass Flow kg/hr 2-BUOH 31.152 17820 48.433 27259.182 5755.356 5451.836 21807.345 17740.427 OLEYLOH 287748.52 287760 trace 14.347 0.003 2.869 11.477 11.4774252 WATER 23737.156 23760 22.7 853.206 24157.721 170.641 682.565 2.98E−07 N—C6 trace 525.936 27471.194 2.87 5494.239 21976.955 0.0177523 CO2 trace 660 660.001 368.622 45.511 73.724 294.897 4.89E−30 Mass Frac 2-BUOH 100 PPM 0.054 0.039 0.487 0.192 0.487 0.487 0.99935245 OLEYLOH 0.924 0.872 trace 256 PPM 94 PPB  256 PPM 256 PPM 0.00064654 WATER 0.076 0.072 0.018 0.015 0.806 0.015 0.015 1.68E−11 N—C6 trace 0.418 0.491 96 PPM 0.491 0.491 1.00E−06 CO2 trace 0.002 0.525 0.007 0.002 0.007 0.007 2.75E−34

TABLE 4 Simulated BUOH Column Traffic and Liquid Mass Composition Profile Outputs for Example 1. Heat Liquid Vapor Liquid Vapor Mixed Liquid Vapor Temperature Pressure duty flow flow feed feed feed product product Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr 1 64.879126 0.986923 0 55627.9935 27021.3176 44762.6688 10.5709356 0 0 27021.3176 2 68.490445 0.990448 0 57674.293 37876.0714 0 0 0 0 0 3 69.235774 0.993973 0 58007.181 39922.3708 0 0 0 0 0 4 69.52182 0.997497 0 57817.0513 40255.2588 0 0 0 0 0 5 70.066442 1.001022 0 56571.8929 40065.1292 0 0 0 0 0 6 72.384545 1.004547 0 52410.0328 38819.9708 0 0 0 0 0 7 80.914387 1.008072 0 48063.3543 34658.111 0 0 0 0 0 8 92.794275 1.011596 0 48126.7796 30311.432 0 0 0 0 0 9 98.252025 1.015121 0 48853.4454 30374.8574 0 0 0 0 0 10 99.729315 1.018646 0 49100.032 31101.5232 0 0 0 0 0 11 100.12846 1.022171 0 49173.6467 31348.1099 0 0 0 0 0 12 100.28719 1.025695 0 49206.62 31421.7245 0 0 0 0 0 13 100.39431 1.02922 0 49230.7751 31454.6978 0 0 0 0 0 14 100.49026 1.032745 0 49252.6876 31478.853 0 0 0 0 0 15 100.58687 1.036269 17366.1216 17751.9221 31500.7655 0 0 0 17751.922 0

TABLE 5 Simulated Solvent Column Traffic and Liquid Mass Composition Profile Outputs for Example 1. Heat Liquid Vapor Liquid Vapor Temperature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 91.603505 0.986923 0 32298.369 59637.9415 41149.9372 4.83408955 2 95.920715 0.993973 0 32111.9083 50781.5393 0 536.382563 3 96.339035 1.001022 0 364047.199 50058.6959 329463.617 0 4 97.277269 1.008072 0 361701.456 52530.3696 0 0 5 98.149382 1.015121 0 359371.797 50184.6253 0 0 6 98.985077 1.022171 0 357409.569 47854.9674 0 0 7 99.712494 1.02922 0 355943.634 45892.7389 0 0 8 100.30332 1.036269 0 354949.471 44426.8042 0 0 9 100.76828 1.043319 0 354326.157 43432.6418 0 0 10 101.13588 1.050368 0 353962.728 42809.3273 0 0 11 101.43569 1.057418 0 353769.003 42445.8987 0 0 12 101.69116 1.064467 0 353681.364 42252.1732 0 0 13 101.91862 1.071517 0 353658.26 42164.5345 0 0 14 102.12867 1.078566 0 353764.202 42141.4305 0 0 15 99.508871 1.085616 92576.5773 311516.83 42247.372 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 59637.9415 18.8581326 32279.511 2 0 0 0 18.0300995 32093.878 3 0 0 0 313405.965 50641.234 4 0 0 0 309238.51 52462.946 5 0 0 0 305174.93 54196.867 6 0 0 0 301724.406 55685.162 7 0 0 0 299111.558 56832.076 8 0 0 0 297300.902 57648.569 9 0 0 0 296124 58202.157 10 0 0 0 295392.773 58569.955 11 0 0 0 294953.299 58815.704 12 0 0 0 294696.585 58984.779 13 0 0 0 294551.291 59106.969 14 0 0 0 294460.276 59303.926 15 0 311516.83 0 0 311516.83

TABLE 6 Simulated Stream Outputs for Example 2. 11 12 19 20 21 30 33 40 41 Temperature C. 79.1 25 79.1 42 64.7 84.1 79.1 32 98.4 Pressure atm 0.99 2 0.99 0.99 0.99 0.99 0.99 1.04 1.09 Vapor Frac 1 0 0.117 0.201 1 1 0.994 0 0 Mole Flow kmol/hr 315.821 324.932 1004.081 1329.022 353.182 966.72 1319.902 1338.63 1068.954 Mass Flow kg/hr 27214.93 28000 41597.72 69598.092 26807.06 42005.59 68812.653 131500 112291.57 Volume Flow l/hr 8.90E+06 42683.8 3.43E+06 6.79E+06 9.56E+06 2.82E+07 3.76E+07 156158.8 139295.86 Enthalpy MMBtu/hr −47.593 −61.151 −286.381 −347.534 −64.982 −233.71 −298.692 −466.49 −361.339 Mass Flow kg/hr 2-BUOH 29985.11 29985.365 4029.378 25955.73 29985.109 18199.6 11.227 OLEYLOH 0.011 0.011 trace 0.011 0.011 99716.45 99716.441 WATER 10238.81 10238.917 710.699 9528.112 10238.81 12597.7 12563.897 N—C6 27214.93 28000 28000 21771.05 5443.883 27214.929 trace CO2 1373.794 1373.8 295.941 1077.852 1373.794 986.25 trace Mass Frac 2-BUOH 0.721 0.431 0.15 0.618 0.436 0.138 100 PPM OLEYLOH 259 PPB 154 PPB trace 256 PPB 157 PPB 0.758 0.888 WATER 0.246 0.147 0.027 0.227 0.149 0.096 0.112 N—C6 1 1 0.402 0.812 0.13 0.395 trace CO2 0.033 0.02 0.011 0.026 0.02 0.007 trace 42 43 46 47 48 49 50 60 Temperature C. 42 76.8 35 35 35 35 35 100.5818 Pressure atm 1.09 1.05 1 1 1 1 1 1.036269 Vapor Frac 0 0.021 1 0 0 0 0 0 Mole Flow kmol/hr 1068.954 1338.63 34.379 746.999 547.644 149.4 597.6 244.4176 Mass Flow kg/hr 112291.57 131500 1877.632 56154.133 11566.33 11230.826 44923.306 18116.24 Volume Flow l/hr 132655.75 938664.6 856669 78481.049 12167.74 15696.21 62784.84 25035.05 Enthalpy MMBtu/hr −375.642 −452.187 −10.484 −195.837 −149.474 −39.167 −156.669 −75.1987 Mass Flow kg/hr 2-BUOH 11.227 18199.6 72.413 27681.992 2230.96 5536.398 22145.594 18116.22 OLEYLOH 99716.441 99716.45 trace 0.011 trace 0.002 0.009 0.008586 WATER 12563.897 12597.7 33.91 888.373 9316.634 177.675 710.699 4.06E−07 N—C6 trace 785.053 27213.83 1.117 5442.766 21771.064 0.018124 CO2 trace 986.25 986.256 369.927 17.617 73.985 295.941 5.30E−30 Mass Frac 2-BUOH 100 PPM 0.138 0.039 0.493 0.193 0.493 0.493 0.999999 OLEYLOH 0.888 0.758 trace 191 PPB trace 191 PPB 191 PPB 4.74E−07 WATER 0.112 0.096 0.018 0.016 0.805 0.016 0.016 2.24E−11 N—C6 trace 0.418 0.485 97 PPM 0.485 0.485 1.00E−06 CO2 trace 0.007 0.525 0.007 0.002 0.007 0.007 2.93E−34

TABLE 7 Simulated BUOH Column Traffic and Liquid Mass Composition Profile Outputs for Example 2. Heat Liquid Vapor Liquid Vapor Mixed Liquid Vapor Temperature Pressure duty flow flow feed feed feed product product Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr kg/hr 1 64.747241 0.986923 0 55801.169 26807.06 44912.7 10.603188 0 0 26807.064 2 68.4750126 0.990448 0 57949.341 37684.93 0 0 0 0 0 3 69.2761477 0.993973 0 58315.738 39833.1 0 0 0 0 0 4 69.5802816 0.997497 0 58116.921 40199.5 0 0 0 0 0 5 70.1577473 1.001022 0 56810.98 40000.68 0 0 0 0 0 6 72.6085042 1.004547 0 52581.194 38694.74 0 0 0 0 0 7 81.3555634 1.008072 0 48407.827 34464.95 0 0 0 0 0 8 93.0514383 1.011596 0 48563.256 30291.59 0 0 0 0 0 9 98.3137126 1.015121 0 49278.822 30447.01 0 0 0 0 0 10 99.7393512 1.018646 0 49519.562 31162.58 0 0 0 0 0 11 100.128653 1.022171 0 49592.207 31403.32 0 0 0 0 0 12 100.285692 1.025695 0 49625.234 31475.96 0 0 0 0 0 13 100.392535 1.02922 0 49649.612 31508.99 0 0 0 0 0 14 100.488452 1.032745 0 49672.109 31533.37 0 0 0 0 0 15 100.581838 1.036269 17396.32 18116.242 31555.87 0 0 0 18116.242 0

TABLE 8 Simulated Solvent Column Traffic and Liquid Mass Composition Profile Outputs for Example 2. Heat Liquid Vapor Liquid Vapor Temperature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 84.060187 0.986923 0 16765.096 42005.59 22792.45 4.7038857 2 89.8926777 0.993973 0 14441.477 35973.53 0 1221.9044 3 92.0868107 1.001022 0 148478.34 32428.01 130278.1 0 4 92.5377696 1.008072 0 148157.13 36186.77 0 0 5 92.9685493 1.015121 0 147191.84 35865.58 0 0 6 93.7304288 1.022171 0 145171.65 34900.31 0 0 7 95.0675997 1.02922 0 142024.42 32880.1 0 0 8 96.9749834 1.036269 0 138714.5 29732.83 0 0 9 98.9018001 1.043319 0 136370.53 26422.91 0 0 10 100.289731 1.050368 0 135132.87 24078.96 0 0 11 101.108403 1.057418 0 134585.28 22841.31 0 0 12 101.581171 1.064467 0 134370.36 22293.74 0 0 13 101.888191 1.071517 0 134299.02 22078.82 0 0 14 102.123286 1.078566 0 134352.04 22007.48 0 0 15 98.4221195 1.085616 48299.74 112291.57 22060.48 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 42005.5887 3195.9415 13569.15 2 0 0 0 10.7623383 14430.71 3 0 0 0 126144.955 22333.38 4 0 0 0 125506.149 22650.98 5 0 0 0 123866.355 23325.49 6 0 0 0 120493.419 24678.23 7 0 0 0 115265.898 26758.52 8 0 0 0 109749.425 28965.07 9 0 0 0 105795.42 30575.11 10 0 0 0 103667.523 31465.35 11 0 0 0 102696.62 31888.66 12 0 0 0 102289.059 32081.3 13 0 0 0 102125.941 32173.08 14 0 0 0 102054.488 32297.55 15 0 112291.565 0 0 112291.6

TABLE 9 Simulated Stream Outputs for Example 3. 11 12 19 20 21 30 33 Temperature C. 90.4 25 90.4 83.9 57.8 93.1 90.4 Pressure atm 0.99 2 0.99 0.99 0.99 0.99 0.99 Vapor Frac 1 0 0.666 0.607 1 1 0.99 Mole Flow kmol/hr 317.903 324.932 2731.109 3056.045 384.357 2664.655 3049.012 Mass Flow kg/hr 27394.32 28000 76093.335 104093.16 26207.542 77280.113 103487.66 Volume Flow l/hr 9.28E+06 42683.8 5.43E+07 5.41E+07 1.02E+07 8.01E+07 9.00E+07 Enthalpy MMBtu/hr −47.335 −61.151 −671.178 −732.33 −70.469 −613.775 −684.244 Mass Flow kg/hr ETOH 0 0 7313.2724 7312.9414 979.86415 6333.4082 7313.2724 I-BUOH 0 0 28581.499 28581.481 1689.573 26891.926 28581.499 OLEYLOH 0 0 23.900627 23.900944 4.00E−05 23.900587 23.900627 WATER 0 0 38791.702 38791.8 1274.3438 37517.359 38791.702 N—C6 27395.64 28000 0 28000 21909.084 5486.5543 27395.638 CO2 0 0 1140.896 1140.8928 338.77087 802.12515 1140.896 Mass Frac ETOH 0 0 0.096416 0.0704176 0.0374113 0.0821931 0.0708328 I-BUOH 0 0 0.3768098 0.2752162 0.0645081 0.3489953 0.2768267 OLEYLOH 0 0 0.0003151 0.0002301 1.53E−09 0.0003102 0.0002315 WATER 0 0 0.511418 0.3735332 0.0486546 0.4868889 0.3757178 N—C6 1 1 0 0.269617 0.8364916 0.0712028 0.265341 CO2 0 0 0.0150412 0.0109859 0.0129343 0.0104097 0.0110502 40 41 42 43 46 47 Temperature C. 32 99.5 42 85.8 35 35 Pressure atm 1.04 1.09 1.09 1.05 1 1 Vapor Frac 0 0 0 0.003 1 0 Mole Flow kmol/hr 2677.57 2382.649 2382.649 2677.57 24.033 827.846 Mass Flow kg/hr 330000 309735.02 309735.02 330000 1330.4 57184.545 Volume Flow l/hr 391006.5 386948.5 368063.78 600454.04 598279.2 79969.918 Enthalpy MMBtu/hr −991.655 −854.235 −893.255 −952.636 −7.161 −211.874 Mass Flow kg/hr ETOH 1781.287 0.0038104 0.0038104 1781.2875 18.3884 3428.0356 I-BUOH 17812.87 30.957457 30.957457 17812.875 22.66507 24311.009 OLEYLOH 285995.6 285976.49 285976.49 285995.6 1.84E−07 23.895623 WATER 23750.5 23567.137 23567.137 23750.5 23.3043 1793.1255 N—C6 0 6.28E−20 6.28E−20 0 604.342 27386.38 CO2 659.7361 3.18E−26 3.18E−26 659.73611 659.7329 423.46358 Mass Frac ETOH 0.005398 1.23E−08 1.23E−08 0.0053978 0.013842 0.0597574 I-BUOH 0.053978 0.0001 0.0001 0.0539784 0.017062 0.4237884 OLEYLOH 0.866653 0.9237725 0.9237725 0.8666533 1.39E−10 0.0004165 WATER 0.071971 0.0761275 0.0761275 0.0719712 0.017543 0.0312577 N—C6 0 2.03E−25 2.03E−25 0 0.454929 0.4773982 CO2 0.001999 1.03E−31 1.03E−31 0.0019992 0.496625 0.0073818 48 49 50 60 Temperature C. 35 35 35 102.5 Pressure atm 1 1 1 1.04 Vapor Frac 0 0 0 0 Mole Flow kmol/hr 2204.166 165.569 662.277 277.92 Mass Flow kg/hr 45578.22 11436.909 45747.636 19540.094 Volume Flow l/hr 48033.624 15993.984 63975.935 26906.753 Enthalpy MMBtu/hr −597.126 −42.375 −169.499 −81.768 Mass Flow kg/hr ETOH 3866.5175 685.607115 2742.4285 1762.5643 I-BUOH 4247.8066 4862.20184 19448.807 17759.234 OLEYLOH 0.0053202 4.77912468 19.116499 19.116459 WATER 36975.371 358.625109 1434.5004 160.15661 N—C6 9.2783454 5477.27593 21909.104 0.0197013 CO2 57.696326 84.6927162 338.77087 2.21E−23 Mass Frac ETOH 0.0856245 0.05975736 0.0597574 0.0894653 I-BUOH 0.0940682 0.42378844 0.4237884 0.901434 OLEYLOH 1.18E−07 0.00041654 0.0004165 0.0009703 WATER 0.818824 0.03125768 0.0312577 0.0081293 N—C6 0.0002055 0.47739816 0.4773982 1.00E−06 CO2 0.0012777 0.00738179 0.0073818 1.12E−27

TABLE 10 Simulated BUOH Column Traffic and Liquid Mass Composition Profile Outputs for Example 3. Temper- Heat Liquid Vapor Liquid Vapor ature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 59.36128 0.986923 0 51520.418 26191.636 45834.123 58.603933 2 60.12645 0.990448 0 51713.885 31819.326 0 0 3 60.25385 0.993973 0 51784.246 32012.794 0 0 4 60.2306 0.997497 0 51698.837 32083.155 0 0 5 60.43466 1.001022 0 51796.771 31997.745 0 0 6 60.39883 1.004547 0 51776.559 32095.679 0 0 7 60.4211 1.008072 0 51765.873 32075.467 0 0 8 60.47596 1.011596 0 49214.802 32064.781 0 0 9 66.54937 1.015121 0 39332.283 29513.71 0 0 10 83.26737 1.018646 0 38067.089 19631.191 0 0 11 88.24657 1.022171 0 38643.249 18365.997 0 0 12 89.60918 1.025695 0 39648.378 18942.158 0 0 13 91.48175 1.02922 0 41563.434 19947.287 0 0 14 95.05938 1.032745 0 44591.651 21862.343 0 0 15 100.302 1.036269 17882.824 19701.091 24890.56 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 26191.636 43926.473 7593.945 2 0 0 0 50451.673 1262.212 3 0 0 0 50498.203 1286.044 4 0 0 0 44052.646 7646.19 5 0 0 0 50517.359 1279.412 6 0 0 0 44098.498 7678.061 7 0 0 0 44001.078 7764.795 8 0 0 0 40958.848 8255.954 9 0 0 0 39332.283 0 10 0 0 0 38067.089 0 11 0 0 0 38643.249 0 12 0 0 0 39648.378 0 13 0 0 0 41563.434 0 14 0 0 0 44591.651 0 15 0 19701.091 0 19701.091 0

TABLE 11 Simulated Solvent Column Traffic and Liquid Mass Composition Profile Outputs for Example 3. Heat Liquid Vapor Liquid Vapor Temperature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 93.1653 0.986923 0 48893.429 77055.272 56615.205 14.650976 2 96.67048 0.993973 0 48356.078 69318.845 0 484.74926 3 97.13599 1.001022 0 379319.1 68296.745 329515.25 0 4 98.11575 1.008072 0 376501.6 69744.519 0 0 5 98.90471 1.015121 0 374043.88 66927.019 0 0 6 99.57766 1.022171 0 372146.27 64469.296 0 0 7 100.1302 1.02922 0 370805.65 62571.687 0 0 8 100.5749 1.036269 0 369924.07 61231.063 0 0 9 100.9347 1.043319 0 369379.04 60349.488 0 0 10 101.2333 1.050368 0 369063.07 59804.455 0 0 11 101.4906 1.057418 0 368895.83 59488.483 0 0 12 101.7208 1.064467 0 368822.36 59321.251 0 0 13 101.9337 1.071517 0 368806.83 59247.773 0 0 14 102.1355 1.078566 0 368953.75 59232.242 0 0 15 99.51284 1.085616 130955.66 309574.58 59379.17 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 77055.272 26.293919 48867.13 2 0 0 0 26.174428 48329.9 3 0 0 0 312017.73 67301.37 4 0 0 0 306968.52 69533.08 5 0 0 0 302652.42 71391.46 6 0 0 0 299295.74 72850.53 7 0 0 0 296884.16 73921.49 8 0 0 0 295254.49 74669.58 9 0 0 0 294201.36 75177.68 10 0 0 0 293542.54 75520.53 11 0 0 0 293140.68 75755.15 12 0 0 0 292901.21 75921.15 13 0 0 0 292762.53 76044.3 14 0 0 0 292667.93 76285.83 15 0 309574.58 0 0 309574.6

TABLE 12 Simulated Stream Outputs for Example 4. 11 12 19 20 21 30 33 40 41 Temperature C. 88.8 25 88.8 82.8 58.9 91.5 88.8 32 99.5 Pressure atm 0.99 2 0.99 0.99 0.99 0.99 0.99 1.04 1.09 Vapor Frac 1 0 0.712 0.648 1 1 0.991 0 0 Mole Flow kmol/hr 317.756 324.932 2728.966 3053.887 378.959 2667.763 3046.721 2698.119 2343.146 Mass Flow kg/hr 27381.63 28000 80650.799 108650.32 26333.043 81699.385 108032.43 330000 307804.64 Volume Flow l/hr 9.23E+06 42683.8 5.77E+07 5.75E+07 1.01E+07 7.98E+07 8.96E+07 391154.47 384657.78 Enthalpy MMBtu/hr −47.395 −61.151 −666.006 −727.155 −69.229 −614.595 −683.824 −995.721 −842.703 Mass Flow kg/hr ETOH 13027.915 13027.535 1427.108 11600.807 13027.915 3284.882 0.007 I-BUOH 30174.766 30174.683 1477.034 28697.732 30174.766 17731.273 30.779 OLEYLOH 18.697 18.697 <0.001 18.697 18.697 284685.43 284670.51 WATER 36252.825 36252.805 1189.746 35063.078 36252.825 23641.697 23103.345 N—C6 27381.63 28000 28000 21885.117 5496.512 27381.629 trace CO2 1176.597 1176.604 354.038 822.559 1176.597 656.714 trace Mass Frac ETOH 0.162 0.12 0.054 0.142 0.121 0.01  22 PPB I-BUOH 0.374 0.278 0.056 0.351 0.279 0.054 100 PPM OLEYLOH 232 PPM 172 PPM 1 PPB 229 PPM 173 PPM 0.863 0.925 WATER 0.45 0.334 0.045 0.429 0.336 0.072 0.075 N—C6 1 1 0.258 0.831 0.067 0.253 trace CO2 0.015 0.011 0.013 0.01 0.011 0.002 trace 42 43 46 47 48 49 50 60 Temperature C. 42 85.4 35 35 35 35 35 95 Pressure atm 1.09 1.05 1 1 1 1 1 1.04 Vapor Frac 0 0.003 0 0 0 0 <0.001 Mole Flow kmol/hr 2343.146 2698.119 24.263 896.043 2133.581 179.209 716.835 337.876 Mass Flow kg/hr 307804.64 330000 1346.45 59749.806 47554.068 11949.961 47799.845 21466.802 Volume Flow l/hr 365863.9 602048.47 83180.38 51387.885 16636.076 66544.304 29143.379 Enthalpy MMBtu/hr −881.463 −956.96 −229.936 −578.686 −45.987 −183.949 −97.749 Mass Flow kg/hr ETOH 0.007 3284.882 28.134 5854.338 7145.064 1170.868 4683.47 3256.362 I-BUOH 30.779 17731.273 20.141 23946.63 6207.912 4789.326 19157.304 17680.269 OLEYLOH 284670.51 284685.43 trace 18.652 0.045 3.73 14.922 14.922 WATER 23103.345 23641.697 23.105 2131.216 34098.483 426.243 1704.973 515.227 N—C6 trace 618.35 27356.423 25.227 5471.285 21885.138 0.021 CO2 trace 656.714 656.721 442.547 77.336 88.509 354.038 trace Mass Frac ETOH  22 PPB 0.01 0.021 0.098 0.15 0.098 0.098 0.152 I-BUOH 100 PPM 0.054 0.015 0.401 0.131 0.401 0.401 0.824 OLEYLOH 0.925 0.863 trace 312 PPM 947 PPB  312 PPM 312 PPM 695 PPM WATER 0.075 0.072 0.017 0.036 0.717 0.036 0.036 0.024 N—C6 trace 0.459 0.458 530 PPM 0.458 0.458 1000 PPB   CO2 trace 0.002 0.488 0.007 0.002 0.007 0.007 trace

TABLE 13 Simulated BUOH Column Traffic and Liquid Mass Composition Profile Outputs for Example 4. Temper- Heat Liquid Vapor Liquid Vapor ature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 58.8677 0.986923 0 53619.022 26333.043 47724.677 75.167578 2 59.54895 0.990448 0 53793.603 32152.221 0 0 3 59.65739 0.993973 0 53804.133 32326.802 0 0 4 59.75322 0.997497 0 53811.405 32337.332 0 0 5 59.84637 1.001022 0 53818.989 32344.603 0 0 6 59.93224 1.004547 0 53827.699 32352.188 0 0 7 59.99611 1.008072 0 53834.924 32360.898 0 0 8 60.02964 1.011596 0 53227.389 32368.122 0 0 9 61.89011 1.015121 0 43249.936 31760.587 0 0 10 78.22766 1.018646 0 39494.953 21783.134 0 0 11 86.2438 1.022171 0 39676.2 18028.151 0 0 12 87.59943 1.025695 0 40017.329 18209.398 0 0 13 88.42887 1.02922 0 40932.433 18550.528 0 0 14 90.35375 1.032745 0 43270.946 19465.631 0 0 15 94.95595 1.036269 17905.467 21466.802 21804.144 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 26333.043 41870.422 11748.6 2 0 0 0 41980.297 11813.306 3 0 0 0 41974.264 11829.869 4 0 0 0 41964.857 11846.548 5 0 0 0 41952.005 11866.984 6 0 0 0 41928.534 11899.165 7 0 0 0 41880.218 11954.706 8 0 0 0 41103.791 12123.598 9 0 0 0 43249.936 0 10 0 0 0 39494.953 0 11 0 0 0 39676.2 0 12 0 0 0 40017.329 0 13 0 0 0 40932.433 0 14 0 0 0 43270.946 0 15 0 21466.802 0 21466.802 0

TABLE 14 Simulated Solvent Column Traffic and Liquid Mass Composition Profile Outputs for Example 4. Temper- Heat Liquid Vapor Liquid Vapor ature Pressure duty flow flow feed feed Stage C. atm MJ/hr kg/hr kg/hr kg/hr kg/hr 1 91.49881 0.986923 0 48598.116 81699.385 59485.237 18.791886 2 96.14796 0.993973 0 47557.727 70793.473 0 487.25557 3 96.78605 1.001022 0 378241.44 69265.827 329512.74 0 4 97.91872 1.008072 0 375236.53 70436.828 0 0 5 98.79444 1.015121 0 372628.96 67431.944 0 0 6 99.51716 1.022171 0 370622.9 64824.22 0 0 7 100.0975 1.02922 0 369210.81 62818.008 0 0 8 100.558 1.036269 0 368286.26 61405.939 0 0 9 100.9262 1.043319 0 367717.06 60481.46 0 0 10 101.2293 1.050368 0 367388.2 59912.283 0 0 11 101.4888 1.057418 0 367214.33 59583.436 0 0 12 101.7201 1.064467 0 367137.63 59409.628 0 0 13 101.9335 1.071517 0 367120.55 59332.935 0 0 14 102.1354 1.078566 0 367265.35 59315.849 0 0 15 99.54878 1.085616 131183.53 307804.64 59460.559 0 0 Mixed Liquid Vapor feed product product 1st liquid 2nd liquid Stage kg/hr kg/hr kg/hr flow kg/hr flow kg/hr 1 0 0 81699.385 29.815222 48568.301 2 0 0 0 29.22518 47528.501 3 0 0 0 311969.16 66272.288 4 0 0 0 306458.19 68778.339 5 0 0 0 301822.64 70806.313 6 0 0 0 298254 72368.896 7 0 0 0 295709.6 73501.208 8 0 0 0 294002.43 74283.833 9 0 0 0 292906.65 74810.408 10 0 0 0 292225.62 75162.582 11 0 0 0 291812.73 75401.603 12 0 0 0 291568.16 75569.473 13 0 0 0 291427.41 75693.14 14 0 0 0 291332.35 75932.999 15 0 307804.64 0 0 307804.64 

What is claimed is:
 1. A process comprising: a) introducing a feed comprising: (i) a water-immiscible organic extractant; (ii) water; (iii) at least one isomer of butanol; (iv) optionally a non-condensable gas; (v) ethanol; and (vi) optionally fermentation by-products into a first distillation column, wherein the first distillation column comprises a stripping section and optionally a rectifying section at an introduction point above the stripping section, the first distillation column having an operating temperature T₁ and an operating pressure P₁ in the stripping section; wherein T₁ and P₁ are selected to produce a first bottoms stream and a first vaporous overhead stream, the first bottoms stream comprising the water-immiscible organic extractant and water and being substantially free of butanol, and the first vaporous overhead stream comprising water, butanol, ethanol, optionally the extractant, and optionally the non-condensable gas; b) introducing a water-immiscible organic entrainer to at least one process stream or vessel; c) condensing the first vaporous overhead stream to produce a gas phase and recover a first mixed condensate; wherein the first mixed condensate comprises: (i) an organic phase comprising butanol, ethanol, entrainer, and water; and (ii) an aqueous phase comprising water, butanol, and ethanol; and wherein the first mixed condensate comprises sufficient entrainer to provide phase separation of the organic and the aqueous phases; d) introducing at least a portion of the aqueous phase to the first distillation column; e) introducing a first portion of the organic phase into a second distillation column having at least a stripping section; and f) operating the second distillation column to produce a second bottoms stream comprising butanol and ethanol and wherein said bottoms stream is substantially free of water and entrainer, and a second vaporous overhead stream comprising butanol, ethanol, entrainer, and water.
 2. The process of claim 1, wherein the extractant is selected such that it (A) preferentially extracts butanol over water and (B) is separable from butanol by distillation; and the entrainer is selected such that it (C) has a higher vapor pressure than butanol and (D) is separable from butanol by distillation.
 3. The process of claim 1, further comprising any one of steps: g) introducing a second portion of the organic phase into the first distillation column; or h) condensing the vaporous overhead stream from the second distillation column to recover a second condensate stream, and introducing at least a portion of the second condensate stream into the first mixed condensate; or j) both steps g) and step h).
 4. The process of claim 1, wherein the process stream or vessel of step (b) is selected from the feed stream, the first vaporous overhead stream, the first mixed condensate stream, the second condensate stream, the portion of the organic phase which is introduced into the second column, the vaporous overhead stream from the second distillation column, the first distillation column, the decanter, the upper section of the second distillation column, and combinations thereof.
 5. The process of claim 1, further comprising the step of introducing water, steam, or a mixture thereof into the first distillation column.
 6. The process of claim 1, further comprising the step of introducing at least a portion of the aqueous phase into a fermentation vessel.
 7. The process of claim 1, wherein the process further comprises the step of introducing the bottoms stream from the first distillation column into a fermentation vessel.
 8. The process of claim 1, wherein the process further comprises separating the bottoms stream from the first distillation column to obtain a bottoms aqueous phase comprising water and a bottoms organic phase comprising the extractant, introducing at least a portion of the bottoms organic phase into a fermentation vessel, and optionally introducing at least a portion of the bottoms aqueous phase into a fermentation vessel.
 9. The process of claim 1, wherein the gas phase comprises entrainer, and the process further comprises condensing the gas phase to recover at least a portion of the entrainer and optionally introducing the recovered entrainer to the first mixed condensate.
 10. The process of claim 9, wherein a non-condensable gas is present in the feed, the gas phase further comprises the non-condensable gas, and the process further comprises purging at least a portion of the gas phase comprising the non-condensable gas from the process.
 11. The process of claim 1, wherein the non-condensable gas comprises carbon dioxide, nitrogen, hydrogen, noble gases, or mixtures.
 12. The process of claim 1, wherein the feed comprises an organic phase obtained from an extractive fermentation.
 13. The process of claim 1, wherein the butanol concentration in the feed is from about 0.1 weight percent to about 40 weight percent, based on the weight of the feed.
 14. The process of claim 1, wherein the entrainer comprises at least one hydrocarbon.
 15. The process of claim 14, wherein said at least one hydrocarbon selected from the group consisting of pentane, hexane, hexene, cyclohexane, benzene, toluene, and xylene.
 16. The process of claim 1, wherein the extractant comprises at least one solvent selected from the group consisting of C₇ to C₂₂ fatty alcohols, C₇ to C₂₂ fatty acids, esters of C₇ to C₂₂ fatty acids, C₇ to C₂₂ fatty aldehydes, C₇ to C₂₂ fatty amides, and mixtures thereof.
 17. The process of claim 16, wherein the extractant comprises at least one solvent selected from the group consisting of C₁₂ to C₂₂ fatty alcohols, C₁₂ to C₂₂ fatty acids, esters of C₁₂ to C₂₂ fatty acids, C₁₂ to C₂₂ fatty aldehydes, C₁₂ to C₂₂ fatty amides, and mixtures thereof.
 18. The process of claim 1, wherein the isomer of butanol is 1-butanol, 2-butanol, isobutanol, or mixtures thereof.
 19. The process of claim 1, wherein the isomer of butanol is isobutanol.
 20. The process of claim 1 wherein the butanol in the feed introduced into a first distillation column is produced by fermentation of corn or sugar cane.
 21. The process of claim 1 wherein there is a process to process heat exchange between the feed introduced to the first distillation column and the first bottoms stream.
 22. The process of claim 1, wherein the fermentation by-products comprise organic acids.
 23. The process of claim 1, wherein the fermentation by-products comprise volatile fermentation by-products.
 24. The process of claim 23, wherein the second vaporous overhead stream comprises the volatile fermentation by-products.
 25. The process of claim 24, wherein at least a portion of the second vaporous overhead stream is purged to remove the volatile fermentation by-products.
 26. The process of claim 1, wherein the feed is obtained from a fermentation vessel or an extractor. 